This invention relates to a continuous process for the production of hydrogen-rich gas. More specifically, the present invention pertains to a method and apparatus for preventing over-reduction of iron oxide-based shift catalyst thereby limiting structural deterioration of the catalyst.
Hydrogen gas may be obtained by the catalytic reaction of carbon monoxide and steam. This reaction is exothermic and is commonly referred to as the water-gas shift reaction or shift reaction: CO+H2O→CO2+H2. The reaction is affected by passing carbon monoxide and water through a bed of a suitable catalyst. The feed gas containing carbon monoxide and water may originate from a steam methane reformer (SMR), autothermal reformer (ATR), partial oxidation (POX) reactor, catalytic partial oxidation (CPO) reactor, coal or other solid feed stock gasifier, or other suitable device known in the art.
A typical conventional water-gas shift catalyst is an iron oxide promoted by chromium oxide (CrO2). The general class of oxides of iron will be referred to herein as iron oxide. This catalyst is referred to commonly as a high temperature shift catalyst because it starts to become active at temperatures in the range of about 600 to 710° F., whereas other shift reaction promoting catalysts operate at lower temperatures. The effluent gas stream leaving a reaction zone containing high temperature shift catalyst is at a temperature in the range of about 715 to 1000° F.
Chromium oxide promoted iron oxide-based shift catalyst is relatively low priced, readily available, and its strength is high at the high temperatures which exist at the exit of the catalyst bed. However, a serious disadvantage is that the reaction rate of iron oxide catalysts at low temperatures is comparatively slow. Accordingly, the inlet temperature of the reactants must be at a minimum of about 600° F.
Limitations of HTS catalyst include high temperature and over-reduction, depending on the feed to the HTS reactor, which is normally the syngas stream produced in a hydrogen/sygas production stage. The hydrogen/syngas production stage is where the carbon containing feedstock is converted into hydrogen/syngas by SMR, ATR, POX reactor, CPO reactor, coal or other solid feed stock gasifier, or other suitable device known in the art. The hydrogen/syngas production stage is generally operated at a pressure in the range 5 to 50 bar abs., and normally in the range 10 to 40 bar abs. The temperature at which the hydrogen/syngas production stage is affected will normally be in the range 700 to 1200° C., particularly 750 to 1100° C.
The temperature rise across the shift reactor is generally a limitation affecting steam consumption in partial oxidation (POX) and gasification-based hydrogen production processes. The CO content in the syngas from a POX unit or a gasification unit is high, typically greater than 40 volume %. The water gas shift reaction is used to convert CO, in the presence of H2O, to the desired product H2 and byproduct CO2, which is removed by a downstream separation process. Since the shift reaction is exothermic, conversion of large amounts of CO in the syngas from a POX unit or a gasification unit releases a large amount of heat, causing large temperature rise across the shift reactor, which leads to catalyst deactivation by sintering.
A conventional method for overcoming this temperature issue in the POX- or gasification-based hydrogen process is to use a series of stages of adiabatic shift reactors with inter-stage cooling, either by heat exchangers or direct quench using liquid water (cf U.S. Pat. Nos. 3,595,619 and 6,409,974). The steam requirement in the shift feed is relatively high (e.g., steam-to-dry gas volume ratio of about 2). The sensible heat of the excess steam is needed to moderate the temperature rise across the shift reactor. This excess use of steam, however, reduces the the thermal efficiency of the process. Rao et al. (PCT application US2004/000926) suggest a configuration and method to address the temperature issue for POX- and gasification-based hydrogen processes.
In contrast, the temperature rise across the shift reactor is small in a catalytic steam reformer-based hydrogen production processes because of lower CO content (e.g. typically less than 10 vol. %) and high H2 content (e.g. typically about 50 vol. %) in the syngas. Accordingly, the temperature rise can be tolerated by a simple one-stage, adiabatic shift reactor. In the conventional catalytic steam reforming process, the steam-to-dry gas ratio in a HTS reactor feed is typically around 0.5, which is much smaller than that in the POX- or gasification-based processes (e.g. steam-to-dry gas volume ratio of about 2). For catalytic steam reforming, the steam-to-dry gas ratio is generally set by the HTS catalyst over-reduction limit, not the temperature rise across the shift reactor.
The HTS catalyst comes from the supplier as hematite (Fe2O3) and is reduced in situ to the active magnetite state (Fe3O4). If the catalyst is reduced further to wustite (FeO) or completely to iron metal (Fe°), its strength will decrease to a point where it begins to lose its physical integrity. A further problem with over-reduction is that both wustite and iron metal can catalyze the Fischer-Tropsch reaction. This has two effects: first, there is a decrease in hydrogen production and second, there is an increase in undesirable byproducts, both paraffins and higher alcohols and amines.
The key to maintaining the catalyst in the proper state for the water-gas shift reaction, but not the Fischer-Tropsch reaction is to control the reducing/oxidizing potential of the feed gas such that the catalyst remains in the magnetite state and not the wustite or metallic iron state. The feed gas entering the high temperature shift reactor has four constituents that affect this balance, CO, CO2, H2 and H2O. The hydrogen and carbon monoxide will reduce the iron, while the carbon dioxide and steam will oxidize it.
Control of the relative concentrations of CO, CO2, H2 and H2O is difficult for all sources of feed gas to the shift reactor. For example, for feed gas from a catalytic steam reformer, measurement is difficult and the actual composition depends on many variables such as the reforming temperature and pressure and the ratio of hydrogen to carbon to oxygen atoms in the feed gas to the reformer. The latter in turn depends on the hydrocarbon feedstock and the steam-to-carbon ratio to the reformer. The steam-to-carbon ratio (S/C ratio) is defined as the (overall) ratio of the moles of steam to moles of carbon atoms in the hydrocarbons in the feed(s) to the reformer. Additionally, it is hard to know what the limits to prevent over-reduction actually are since the catalyst damage (over-reduction) comes before the symptoms (byproduct formation and increased pressure drop).
Historically, many plants have operated at conditions where the over-reduction of high temperature shift (HTS) catalyst was not an issue. Through the 1970s and into the early 1980s, hydrogen and ammonia plants operated at steam-to-carbon (S/C) ratios of 3.5 and above. Under these conditions, the HTS catalyst remained in the proper state and over reduction of the catalyst was not an issue. Many of these plants needed the steam for reboiler duty in the acid gas removal system. As more PSA based hydrogen plants were designed and more efficient acid gas removal processes for ammonia plants were developed and introduced to the marketplace, the need for low level heat decreased and operators started reducing the S/C ratio to the reformer for economic reasons. As the S/C ratio declined, catalyst manufacturers and operators struggled to define the acceptable operating range of S/C ratio for the high temperature shift.
The carbon monoxide to carbon dioxide molar ratio and the proportion of steam in the feed to the HTS reactor will depend on the conditions employed in the hydrogen/syngas production stage. In the catalytic steam reforming case, increasing the outlet temperature of the reformer, increasing the pressure, and/or decreasing the steam to feedstock carbon ratio (steam-to-carbon ratio) in the reformer feed, all tend to increase the risk of over-reduction of the shift catalyst in the subsequent shift reactor stage.
Generally to minimize risk of over reduction of shift catalyst in a subsequent high temperature shift stage employing an iron oxide catalyst, it has generally been necessary to employ a gas mixture containing a substantial amount of steam (so that the steam to dry gas molar ratio is greater than about 0.5, or greater than 0.6) and/or to employ hydrogen/syngas production conditions such that the molar ratio of carbon monoxide to carbon dioxide in the gas stream is limited to no more than about 1.9, or no more than 1.8, or no more than 1.7.
Where the hydrogen/syngas production process involves catalytic steam reforming, it is possible to operate with a sufficient excess of steam that such problems are avoided. However the generation of such an excess of steam is not energy efficient and, in the interests of economy, it is desirable to operate steam reforming processes at low steam-to-carbon ratios. In fact, the quest to improve the overall economics of catalytic steam reformer produced H2 has already driven the steam-to-carbon ratio below the point where the syngas produced by the catalytic steam reformer is able to maintain the HTS catalyst in the proper oxidation state. In general, the limit on the steam-to-carbon ratio to a catalytic steam reformer below which an HTS catalyst in the shift reactor will become over-reduced by catalytic steam reformer syngas is approximately 2.8. Today, catalytic steam reformer designs may be developed for steam-to-carbon ratios of 2.5 and lower so that the traditional HTS shift reactor can no longer be used without damage to the catalyst.
It is possible to adjust the composition into the HTS reactor by operating the catalytic steam reformer at low steam-to-carbon ratio (i.e. S/C=2.5) and then adding steam to the catalytic steam reformer syngas immediately upstream of the HTS reactor in order to adjust the HTS inlet composition to prevent over-reduction of the HTS catalyst. Unfortunately, the economic benefits that were achieved by lowering the steam-to-carbon ratio into the catalytic steam reformer process are essentially cancelled out due to the efficiency penalty associated with the added steam injected upstream of the HTS reactor. The overall steam-to-carbon ratio (including steam to catalytic steam reformer plus added steam injection to HTS reactor) required to protect the HTS catalyst from overreduction is approximately 2.8.
Alternatively, a different catalyst may be used that is not damaged by the more reducing stream from a catalytic steam reformer operating with a steam-to-carbon ratio less than 2.8.
The current invention solves the problem of over-reduction of iron oxide based shift catalyst.
Related disclosures include U.S. Pat. Nos. 3,595,619, 4,152,407, 4,341,737, 4,861,745, 4,423,022, 5,030,440, and 6,500,403, and PCT application US2004/000926.